Process for producing middle distillates by hydroisomerization and hydrocracking of a heavy fraction derived from a fischer-tropsch effluent

ABSTRACT

The present invention describes a process for producing middle distillates from a C5+ liquid paraffinic fraction, termed a heavy fraction, with an initial boiling point in the range 15° C. to 40° C. produced by Fischer-Tropsch synthesis, comprising the following steps in succession: passing said C5+ liquid paraffinic fraction, termed a heavy fraction, over at least one ion exchange resin at a temperature in the range 80° C. to 150° C., at a total pressure in the range 0.7 to 2.5 MPa, at an hourly space velocity in the range 0.2 to 2.5 h −1 ; eliminating at least a portion of the water formed in step a); hydrogenating the unsaturated olefinic type compounds of at least a portion of the effluent derived from step b) in the presence of hydrogen and a hydrogenation catalyst; and hydroisomerization/hydrocracking of at least a portion of the hydrotreated effluent derived from step c) in the presence of hydrogen and a hydroisomerization/hydrocracking catalyst.

In the Fischer-Tropsch process, synthesis gas (CO+H₂) is catalyticallytransformed into oxygen-containing products and into essentiallystraight-chain hydrocarbons in the gas, liquid or solid form.

The paraffinic feed produced by Fischer-Tropsch synthesis used in theprocess of the invention is produced from a synthesis gas in theFischer-Tropsch process; the synthesis gas (CO+H₂) is advantageouslyproduced employing three routes.

In one preferred implementation, synthesis gas (CO+H₂) is produced fromnatural gas using the gas-to-liquid, GTL, route.

In another preferred implementation, synthesis gas (CO+H₂) is producedfrom coal using the coal-to-liquid process, CTL.

In another preferred implementation, synthesis gas (CO+H₂) is producedfrom biomass using the biomass-to-liquid process, BTL.

However, such products, principally constituted by normal paraffins,cannot be used as they are, in particular because their cold propertiesare not very compatible with the normal uses of oil cuts. As an example,the pour point of a straight-chain hydrocarbon containing 20 carbonatoms per molecule (boiling point equal to approximately 340° C., i.e.usually included in the middle distillates cut) is approximately 37° C.,which renders its use impossible, since the specification for gas oil is−15° C. Thus, hydrocarbons derived from the Fischer-Tropsch processcomprising mainly n-paraffins have to be transformed into moreupgradeable products such as gas oil or kerosene, for example, which areobtained, for example, after catalytic hydrocracking/hydroisomerizationreactions. In contrast, they may have a non-negligible quantity ofunsaturated compounds of the olefinic type and oxygen-containingproducts (such as alcohols, carboxylic acids, ketones, aldehydes andesters). Moreover, such oxygen-containing and unsaturated compounds areconcentrated in the light fractions. Thus, in the C5+ fractioncorresponding to products boiling at an initial boiling point in therange 15° C. to 40° C., these compounds represent in the range 10-20% byweight of unsaturated olefinic type compounds and in the range 5-10% byweight of oxygen-containing compounds.

Such products are generally free of heteroatomic impurities such assulphur or nitrogen, but may contain small quantities of Fe, Co, Zn, Nior Mo originating from the dissolution of catalyst fines by thecarboxylic acids. Those metals may form complexes with theoxygen-containing compounds. Said products contain no or practically noaromatics, naphthenes and more generally cycles, in particular in thecase of cobalt catalysts.

The hydrogenation of unsaturated olefinic type compounds present inhydrocarbons from the Fischer-Tropsch process is a highly exothermicreaction. Thus, under the severe hydrocracking/hydroisomerizationoperating conditions, the transformation of said unsaturated compoundsmay have a negative impact on the hydrocracking step, such as thermalrunaway of the reaction, substantial coking of the catalyst or theformation of gum by oligomerization. In order to protect thehydrocracking step, a hydrotreatment step is carried out underconditions which are less severe than those of hydrocracking step.However, the impurities in the feed, the oxygen-containing compounds andthe metals (Fe, Co, Zn, Ni, Mo) have a deleterious effect not only onthe activity of the hydrotreatment and hydrocracking catalysts, but alsoon the stability of the hydrotreatment catalyst. In fact, in thehydrotreatment reactor, the operating conditions as regards temperatureare such that the oxygen-containing compounds do not decompose but areadsorbed onto the catalyst and form coke. In the hydrocracking section,the severe operating conditions cause the decomposition ofoxygen-containing compounds into water, CO and CO₂ which are inhibitorsof the acid functions (water) and the hydrogenating function (CO, CO₂)of the hydrocracking catalyst and which thus modifies the activity andselectivity. As a consequence, the presence of alcohol or acid typeoxygen-containing compounds present in the feeds necessitates anincrease in the temperature of the hydrotreatment and hydrocracking stepin order to compensate for the drop in activity and maintain theconversion. Further, the carboxylic acids can extract active particlesfrom the hydrotreatment and hydrocracking catalysts, thus reducing theservice life of said catalysts. Similarly, the metals complexed by theoxygen-containing compounds decompose on the active site of saidcatalysts (HDT and HCK) in the presence of hydrogen and very selectivelypoison the active sites of said catalysts.

One of the aims of the invention is thus to reduce the total oxygencontent of the feed and thus to limit the inhibiting effects of theoxygen-containing compounds and thereby limit the increase in thetemperature in order to compensate for the drop in activity and maintainthe conversion on the two steps, hydrotreatment and hydrocracking.

Thus, upstream of the hydrotreatment step and in order to increase theservice life of the hydrotreatment catalyst and the hydrocrackingcatalyst, the Applicant has instigated a step allowing transformation onan ion exchange resin, simultaneously or otherwise, of the alcohols andcarboxylic acids constituting the oxygen-containing compounds intoesters, and of capturing the metals complexed by said oxygen-containingcompounds.

This step is followed by separation of water before the hydrotreatmentstep, which can reduce the total oxygen content and thus limit theinhibiting effects of the oxygen-containing compounds, and thereby limitthe increase in temperature in order to compensate for the drop inactivity and maintain the conversion over the two steps, hydrotreatmentand hydrocracking. The water separation can also wash and capture CO andCO₂, which are inhibitors, dissolved in the feed.

PRIOR ART

Shell's patent application (EP-0 583 836) describes a process forproducing middle distillates from a feed obtained by the Fischer-Tropschprocess. In this process, the feed from the Fischer-Tropsch synthesismay be treated in its entirety, but preferably the C4− fraction isremoved from the feed so that only the C5+ fraction boiling at atemperature of over 15° C. is introduced into the subsequent step. Saidfeed undergoes hydrotreatment in order to hydrogenate the olefins andalcohols in the presence of a large excess of hydrogen, so that theconversion of products boiling above 370° C. into products with a lowerboiling point is less than 20%. The hydrotreated effluent constituted byparaffinic hydrocarbons with a high molecular weight is preferablyseparated from hydrocarbon compounds with a low molecular weight, inparticular the C4− fraction, before the second hydroconversion step. Atleast a portion of the remaining C5+ fraction then undergoes ahydrocracking/hydroisomerization step with at least 40% by weightconversion of products boiling above 370° C. into products with a lowerboiling point.

Neither the presence of impurities in the feed nor the presence of stepsfor eliminating such impurities is mentioned in that application. Thus,Shell's patent application (EP-0 583 836) does not deal with the problemof eliminating the impurities present in the feed from theFischer-Tropsch process.

SASOL's patent applications (WO-06/005084) and WO-06/005085 concern theelimination of metals complexed by the oxygen-containing compoundspresent in a paraffinic feed derived from the Fischer-Tropsch process.Those patents describe decomposition after adding water in a zone forhydrothermal conversion of those compounds. That decomposition isfollowed by a physical treatment which can remove the metals afterdecomposition. Those applications require adding water to the system,the existence of three steps (reaction, water separation, filtration)and do not affect the oxygen-containing compounds present in the feed.The present invention can reduce the number of steps required, dispensewith the addition of water and simultaneously carry out transformationof the oxygen-containing compounds present in the feed.

More precisely, the present invention concerns a process for producingmiddle distillates from a C5+ liquid paraffinic fraction, termed a heavyfraction, with an initial boiling point in the range 15° C. to 40° C.,produced by Fischer-Tropsch synthesis, comprising the following steps insuccession:

-   -   a) passing said C5+ liquid paraffinic fraction, termed a heavy        fraction, over at least one ion exchange resin at a temperature        in the range 80° C. to 150° C., at a total pressure in the range        0.7 to 2.5 MPa, at an hourly space velocity in the range 0.2 to        2.5 h⁻¹;    -   b) eliminating at least a portion of the water formed in step        a);    -   c) hydrogenating the unsaturated olefinic type compounds of at        least a portion of the effluent derived from step b) in the        presence of hydrogen and a hydrogenation catalyst;    -   d) hydroisomerization/hydrocracking of at least a portion of the        hydrotreated effluent derived from step c) in the presence of        hydrogen and a hydroisomerization/hydrocracking catalyst;    -   e) separating and recycling unreacted hydrogen and light gases        to the hydroisomerization/hydrocracking step d);    -   f) distilling the effluent derived from step e).

BRIEF DESCRIPTION OF THE DRAWINGS

FIGS. 1 and 2 show preferred implementations of the process of theinvention without limiting its scope.

DETAILED DESCRIPTION OF THE INVENTION

Throughout the remainder of the description, we shall detail the varioussteps in the process of the invention by referring to FIGS. 1 and 2.

At the outlet from the Fischer-Tropsch synthesis unit, the effluentderived from the Fischer-Tropsch synthesis unit is advantageouslydivided into two fractions, a light fraction termed the cold condensate,and a heavy fraction, termed waxes.

The two fractions defined thereby comprise water, carbon dioxide (CO₂),carbon monoxide (CO) and unreacted hydrogen (H₂). Further, the lightfraction, the cold condensate, contains light C1 to C4 hydrocarbons,termed the C4− fraction, in the gaseous form.

In accordance with a preferred implementation which is not shown in thefigures, the light fraction, termed the cold condensate, and the heavyfraction, termed waxes, are treated separately then recombined in theline in order to obtain a single C5+ fraction 1, termed the heavyfraction, with an initial boiling point in the range 15° C. to 40° C.,preferably with a boiling point of 20° C. or more.

The light fraction, termed the cold condensate, enters a fractionationmeans which is not shown in the figures. The fractionation means may,for example, be constituted by processes which are well known to theskilled person such as a flash drum, a distillation or a stripper.Advantageously, the fractionation means is a distillation column whichcan eliminate light and gaseous C1 to C4 hydrocarbons, termed thegaseous fraction C4−, corresponding to products boiling at a temperatureof less than 20° C., preferably less than 10° C. and more preferablyless than 0° C., and recovery of a C5+ fraction, termed the heavyfraction, with an initial boiling point in the range 15° C. to 40° C.,preferably with a boiling point of 20° C. or more.

The stabilized effluent derived from the fractionation means and saidheavy fraction, termed waxes, are then recombined in order to obtain astabilized C5+ liquid fraction 1, corresponding to products boiling atan initial boiling point in the range 15° C. to 40° C., preferably witha boiling point or 20° C. or more. Said C5+ fraction 1 constitutes thefeed used in the process of the invention.

Before passage over an ion exchange resin in accordance with step a) ofthe process of the invention, said liquid C5+ fraction may optionallyinitially undergo a step for decontamination in a reactor 2 by passageover a guard bed containing at least one guard bed catalyst.

The treated heavy fractions may possibly contain solid particles such asmineral solids. They may possibly contain metals contained inhydrocarbon structures such as organometallic compounds of greater orlesser solubility. The term “fines” means fines resulting from physicalor chemical attrition of the catalyst. They may be on the micron orsub-micron scale. These mineral particles thus contain the activecomponents of these catalysts; a non limiting list is as follows:alumina, silica, titanium, zirconia, cobalt oxide, iron oxide, tungsten,ruthenium oxide, etc. These solid minerals may be in the form of acalcined mixed oxide: examples are alumina-cobalt, alumina-iron,alumina-silica, alumina-zirconia, alumina-titanium,alumina-silica-cobalt, alumina-zirconia-cobalt, etc.

Said heavy fractions may also contain metals within hydrocarbonstructures, which may possibly contain oxygen or organometalliccompounds of greater or lesser solubility. More particularly, saidcompounds may be silicon-based. As an example, they may be anti-foamingagents used in the synthesis process. As an example, the solutions of asilicone type silicon compound or silicone oil emulsion are moreparticularly contained in the heavy fraction.

Further, the catalyst fines described above may have a silica contentwhich is greater than the formulation for the catalyst, resulting fromintimate interaction between the catalyst fines and the anti-foamingagents described above.

The problem which thus arises is to reduce the quantity of solid mineralparticles and possibly to reduce the quantity of metallic compoundswhich are deleterious to the hydroisomerization-hydrocracking catalyst.

Characteristics of Catalysts used in the Guard Beds

The guard beds advantageously contain at least one catalyst.

Shape of Catalysts

The catalysts in the guard beds used in the optional decontaminationstep of the process of the invention may advantageously have the shapeof spheres or extrudates. However, it is advantageous for the catalystto be in the shape of extrudates with a diameter in the range 0.5 to 5mm, more particularly in the range 0.7 to 2.5 mm. The shapes arecylinders (which may or may not be hollow), twisted cylinders,multilobes (2, 3, 4 or 5 lobes, for example), or rings. The cylindricalshape is preferred, but any other shape may be used.

To accommodate the presence of contaminants and/or poisons in the feed,in a further preferred implementation, the guard bed catalysts may havemore particular geometrical forms to increase their void fraction. Thevoid fraction of said catalysts is in the range 0.2 to 0.75. Theirexternal diameter may be between 1 and 35 mm Possible particularnon-limiting forms are: hollow cylinders, hollow rings, Raschig rings,toothed hollow cylinders, crenellated hollow cylinders, pentaringcartwheels, multiple holed cylinders, etc.

Active Phase

Said catalysts of the guard beds used in the optional decontaminationstep of the process of the invention may advantageously have beenimpregnated with a phase which may or may not be active. Preferably, thecatalysts are impregnated with a hydrodehydrogenating phase. Highlypreferably, the CoMo or NiMo phase is used. Still more preferably, theNiMo phase is used.

Preferably, the supports for said guard bed catalysts are porousrefectory oxides, preferably selected from alumina and silica-alumina.

Said guard bed catalysts may advantageously have macroporosity.

Said catalysts advantageously have a macroporous mercury volume for amean diameter of 50 nm which is more than 0.1 cm³/g, preferably in therange 0.125 to 0.175 cm³/g, and a total volume of more than 0.60 cm³/g,preferably in the range 0.625 to 0.8 cm³/g, and is advantageouslyimpregnated with an active phase, preferably based on nickel andmolybdenum, such as ACT961, for example. In this preferred embodiment,the Ni content as the weight of oxide is generally in the range 1% to10% and the Mo content as the weight of oxide is in the range 5% to 15%.The surface areas, expressed as the SBET, of the supports for saidcatalysts are in the range 30 m²/g to 220 m²/g.

In a first embodiment, the guard bed advantageously also comprises atleast one other catalyst having a mercury volume for a pore diameter ofmore than 1 micron of more than 0.2 cm³/g and preferably more than 0.5cm³/g, and a mercury volume for a pore diameter of more than 10 micronsof more than 0.25 cm³/g and preferably less than 0.4 cm³/g, saidcatalyst advantageously being placed upstream of the catalyst of theinvention.

In a second embodiment, the guard bed advantageously also comprises atleast one other catalyst with a mercury volume for a pore diameter ofmore than 50 nm of more than 0.25 cm³/g, the mercury volume for a porediameter of more than 100 nm being more than 0.15 cm³/g and a total porevolume of more than 0.80 cm³/g.

Said guard bed catalyst and the catalyst of the first embodiment mayadvantageously be associated in a mixed bed or a combined bed. Ingeneral, the impregnated catalyst of the active phase constitutes themajority of the guard bed and the catalyst of the first embodiment whichis preferred is added as a complement of 0 to 50% by volume with respectto the first catalyst, preferably 0 to 30%, more preferably 1% to 20%.

The combination of the catalyst of the invention and the catalyst of thefirst embodiment does not limit the scope of the invention. Thecatalysts which can be used in the guard beds may advantageously be usedalone or as a mixture; in a non-exhaustive manner, they may be selectedfrom those sold by Norton-Saint-Gobain, for example MacroTrap® guardbeds, or catalysts sold by Axens from the ACT family: ACT077, ACT935,ACT961 or HMC841, HMC845, HMC941 or HMC945.

Preferred guard beds for use in the invention are the HMCs and ACT961.

It may be particularly advantageous to superimpose these catalysts in atleast two different beds of varying heights. The catalysts with thehighest void ratio are preferably used in the first catalytic bed orbeds at the inlet to the catalytic reactor. It may also be advantageousto use at least two different reactors for said catalysts.

Advantageously, an association of said guard bed catalyst with thecatalysts of the first and second implementation is also possible in amixed bed or a combined bed. In this case, the catalysts are placed withthe void capacity decreasing in the direction of flow.

After passing over said guard bed, the quantity of solid particles isless than 20 ppm, preferably less than 10 ppm and more preferably lessthan 5 ppm. The soluble silicon content is less than 5 ppm, preferablyless than 2 ppm and more preferably less than 1 ppm.

Step a)

In accordance with step a) of the process of the invention, said C5+liquid paraffinic fraction, termed the heavy fraction, with an initialboiling point in the range 15° C. to 40° C. produced by Fischer-Tropschsynthesis passes over at least one ion exchange resin which can esterifythe alcohols and carboxylic acids into esters and/or capture metalsdissolved in the feed, at a temperature in the range 80° C. to 150° C.,at a total pressure in the range 0.7 to 2.5 MPa, and at an hourly spacevelocity in the range 0.2 to 2.5 h⁻¹.

Step a) of the invention may advantageously be carried out in accordancewith two distinct implementations, namely either in a single reactor 4over a single ion exchange resin, advantageously used to simultaneouslycarry out esterification of alcohols and carboxylic acids to esters andcapture of the metals dissolved in the feed, or in two differentreactors on two ion exchange resins 3 and 4 of different natures, onehaving the specific function of esterification of alcohols andcarboxylic acids and the other the capture of the metals dissolved inthe feed.

In a first implementation, step a) advantageously consists of passingsaid C5+ liquid paraffinic fraction over a single ion exchange resin ina single reactor 4 to simultaneously carry out the esterification ofalcohols and carboxylic acids to esters and the capture of metalsdissolved in the feed.

Preferably, said resin is used at a temperature in the range 100° C. to150° C. and preferably in the range 100° C. to 130° C., at a pressure inthe range 1 to 2 MPa and preferably in the range 1 to 1.5 MPa, and at anhourly space velocity in the range 0.5 to 2 h⁻¹, preferably in the range0.5 to 1.5 h⁻¹.

In this case, oxygen-containing compounds, carboxylic acids and alcoholsare adsorbed onto the active sites of said resin and are esterified andthe cationic and metallic compounds present in said C5+ liquidparaffinic fraction are eliminated by adsorption or by ion exchange.Said resin, which can simultaneously carry out the esterification ofalcohols and carboxylic acids to esters and the capture of metalsdissolved in the feed, advantageously comprises sulphonic acid groupsand is prepared by polymerization or co-polymerization of aromatic vinylgroups followed by sulphonation, said aromatic vinyl groups beingselected from styrene, vinyl toluene, vinyl naphthalene, vinyl ethylbenzene, methyl styrene, vinyl chlorobenzene and vinyl xylene, saidresin having a degree of cross-linking in the range 20% to 35%,preferably in the range 25% to 35%, preferably equal to 30%, and an acidstrength, assayed by potentiometry during neutralization with a KOHsolution, of 0.2 to 6 mmol H+ equivalent/g, preferably in the range 0.2to 2.5 mmol H+ equivalent/g.

Said acid ion exchange resin advantageously contains in the range 1 to 2terminal sulphonic groups per aromatic group. Preferably, said resin hasa size in the range 0.15 to 1.5 mm The size of a resin is the diameterof the sphere encompassing the resin particle. The resin size categoriesare measured by screening through suitable screens, in accordance with atechnique which is known in the art.

A preferred resin is a resin constituted by co-polymers of monovinylaromatics and polyvinyl aromatics, and highly preferably, a copolymer ofdivinyl benzene and polystyrene with a degree of cross-linking in therange 20% to 35%, preferably in the range 25% to 35%, and morepreferably equal to 30%, and an acid strength, representing the numberof active sites of said resin, assayed by potentiometry duringneutralization using a KOH solution, in the range 0.2 to 6 mmol H+equivalent/g, preferably in the range 0.2 to 2.5 mmol H+ equivalent/g.

Another preferred resin which can simultaneously allow theesterification of alcohols and carboxylic acids and the capture ofmetals to be carried out is a resin constituted by a polysiloxanegrafted with alkylsulphonic type acid groups (of the —CH₂—CH₂—CH₂—SO₃Htype), with a size in the range 0.5 to 1.2 mm and with an acid strength,representing the number of active sites of said resin and assayed bypotentiometry during neutralization with a KOH solution, of 0.4 to 1.5mmol H+ equivalent/g.

During said step and under these conditions, 95% of the carboxylic acidsare esterified. The acid conversion is analyzed by the potassiumhydroxide titration difference between the feed and the effluent using atechnique which is known to the skilled person. The ASTM methods D 664,D 3242 or D 974 can be cited, for example, as methods for carrying outsaid analysis.

This resin may advantageously be used in a fixed bed between screensplaced in an upflow or downflow tube reactor. Preferably, said resin isused in an upflow bed reactor, the liquid being injected into the bottomof the reactor at a sufficient surface velocity to allow the bed ofresin to expand without, however, either transporting or fluidizing it.This implementation, compared with a fixed bed, can attenuate theeffects of clogging materials and substantially increase the servicelife of the resin.

In accordance with a second implementation, step a) advantageouslyconsists of passing said C5+ liquid paraffinic fraction 1 into twodifferent reactors 3 and 4 shown in FIG. 2, over two distinct ionexchange resins, of different natures, one having the specific functionof esterification of alcohols and carboxylic acids and the other that ofcapturing the metals dissolved in the feed.

Preferably, the reactor 3 containing the ion exchange resin which cancapture metals is used upstream of the reactor 4 containing the ionexchange resin which can carry out the esterification of the alcoholsand carboxylic acids.

In this case, the cationic and metallic compounds present in said C5+liquid paraffinic fraction are eliminated by adsorption or by ionexchange on a first ion exchange resin. This first resin, which isspecific for capturing metals, advantageously comprises sulphonic acidgroups and is advantageously prepared by polymerization orco-polymerization of aromatic vinyl groups followed by sulphonation,said aromatic vinyl groups advantageously being selected from styrene,vinyl toluene, vinyl naphthalene, vinyl ethyl benzene, methyl styrene,vinyl chlorobenzene and vinyl xylene, said resin having a degree ofcross-linking in the range 1% to 20%, preferably in the range 2% to 8%,and an acid strength, representing the number of active sites in saidresin, assayed by potentiometry during neutralization with a KOHsolution, in the range 1 to 15 mmol H+ equivalent/g, preferably in therange 2.5 to 10 mmol H+ equivalent/g.

Said acid ion exchange resin advantageously contains between 1 and 2terminal sulphonic acid groups per aromatic group. Preferably, saidresin has a size in the range 0.15 to 1.5 mm The size of the resin isthe diameter of the sphere encompassing the resin particle. The sizeclasses for the resin are measured by screening through suitable screensusing a technique which is known to the skilled person.

Preferably, the first resin is a resin constituted by copolymers ofmonovinyl aromatics and polyvinyl aromatics; more preferably, acopolymer of divinyl benzene and polystyrene with a degree ofcross-linking in the range 1% to 20% and an acid strength, representingthe number of active sites of said resin and assayed by potentiometryduring neutralization with a KOH solution, of 1 to 15 mmolH+equivalent/g, preferably in the range 2.5 to 10 mmol H+ equivalent/g.

Preferably, said first resin is used at a temperature in the range 80°C. to 110° C., at a pressure in the range 1 to 2 MPa and preferably inthe range 1 to 1.5 MPa, and at an hourly space velocity in the range 0.2to 1.5 h⁻¹, preferably in the range 0.5 to 1.5 h⁻¹.

The effluent from the reactor 3 containing said first resin which isspecific for the capture of metals is then introduced into a secondreactor 4 located downstream of the first reactor and containing asecond resin with a different nature and which is specific toesterification of the alcohols and carboxylic acids contained in saideffluent.

The oxygen-containing compounds, carboxylic acids and alcohols areadsorbed onto the active sites of said second resin and are esterifiedand the cationic and metallic compounds present in the effluent fromreactor 3 are eliminated by adsorption or by ion exchange. Said secondresin, which can simultaneously carry out the esterification of alcoholsand carboxylic acids to esters and the capture of metals dissolved inthe feed, advantageously comprises sulphonic acid groups and isadvantageously prepared by polymerization or copolymerization ofaromatic vinyl groups followed by sulphonation. The aromatic vinylgroups are advantageously selected from styrene, vinyl toluene, vinylnaphthalene, vinyl ethyl benzene, methyl styrene, vinyl chlorobenzeneand vinyl xylene, said second resin having a degree of cross-linking,i.e. a ratio of the mass of copolymer/mass of polymer, which isadvantageously in the range 20% to 35%, preferably in the range 25% to35% and more preferably 30%, and an acid strength, representing thenumber of active sites of said resin, assayed by potentiometry duringneutralization with a KOH solution, in the range 0.2 to 6 mmolH+equivalent/g, preferably in the range 0.2 to 6 mmol H+ equivalent/g.

Said second acid ion exchange resin advantageously contains 1 to 2terminal sulphonic groups per aromatic group. Preferably, said secondresin has a size in the range 0.15 to 1.5 mm.

A preferred second resin is a resin constituted by copolymers ofmonovinyl aromatics and aromatic polyvinyls, and more preferably, acopolymer of divinyl benzene and polystyrene with a degree ofcross-linking in the range 20% to 35%, preferably in the range 25% to35% and more preferably 30%, and an acid strength, representing thenumber of active sites of said resin, assayed by potentiometry duringneutralization with a KOH solution, in the range 0.2 to 6 mmol H+equivalent/g and preferably in the range 0.2 to 6 mmol H+ equivalent/g.

Another preferred resin which can simultaneously allow esterification ofalcohols and carboxylic acids and capture of metals to be carried out isa resin constituted by a polysiloxane grafted with alkylsulphonic typeacid groups (of the —CH₂—CH₂—CH₂—SO₃H type), with a size in the range0.5 to 1.2 mm and with an acid strength, representing the number ofactive sites of said resin and assayed by potentiometry duringneutralization with a KOH solution, of 0.4 to 1.5 mmol H+ equivalent/g.

Preferably, said second resin is used at a temperature in the range 100°C. to 150° C. and preferably in the range 100° C. to 130° C., at apressure in the range 1 to 2 MPa and preferably in the range 1 to 1.5MPa, and at an hourly space velocity in the range 0.5 to 2 h⁻¹,preferably in the range 0.5 to 1.5 ⁻¹.

During this step and under these conditions, 95% of the carboxylic acidsare esterified. Analysis of the conversion of the acids is given by thedifference in the potassium hydroxide titration between the feed and theeffluent. The ASTM methods D 664, D 3242 or D 974 can be cited, forexample, as methods for carrying out said analysis.

These resins may advantageously be used in a fixed bed between screensplaced in an upflow or downflow tube reactor. Preferably, said resin isused in an upflow bed reactor, the liquid being injected into the bottomof the reactor at a sufficient surface velocity to allow the bed ofresin to expand without, however, either transporting or fluidizing it.This implementation, compared with a fixed bed, can attenuate theeffects of clogging materials and substantially increase the servicelife of the resin.

In the case in which said C5+ liquid paraffinic fraction passes into twodifferent reactors over two distinct ion exchange resins of differentnatures, one principally carrying out metals capture, the otherprincipally carrying out esterification, intermediate re-heating ispreferably employed between the two steps. Preferably, the water formedduring the step for esterification of the acids and alcohols over thefirst resin principally carrying out the capture of metals is removed inorder to intensify the esterification reaction over the second resin.Simultaneously adding intermediate re-heating and water separationboosts the overall conversion of the carboxylic acids present in thefeed.

The reaction for esterification of the organic acids by the alcoholspresent in said C5+ liquid paraffinic fraction, termed a heavy fraction,with an initial boiling point in the range 15° C. to 40° C., produced bythe Fischer-Tropsch synthesis, produces water which is a compound thatinhibits the hydrotreatment and hydrocracking catalysts, necessitatingan increase in the severity of the operating conditions.

Step b)

In accordance with the invention, the effluent derived from step a) thenundergoes a step for eliminating at least a portion of the water formedduring said step a), preferably all of the water formed, in a separator6.

This water is acidic in nature as it advantageously contains protonsexchanged during capture of the metals by the upstream specific cationexchange resin or by the only resin allowing simultaneous esterificationand capture of metals. This water may also contain a fraction ofdissolved CO and CO₂ originating from the Fischer-Tropsch synthesis. Thewater is eliminated via the line 7.

It is also advantageous to add in said step b) gas of the nitrogen (N₂)or hydrogen (H₂) type, 5, to eliminate more dissolved CO and CO₂ bystripping.

In the case when hydrogen is added, this advantageously acts as a makeupgas for the hydrotreatment step.

This step can also eliminate products of the light ether type formedduring the reaction of alcohols with themselves.

The water may be eliminated using any of the methods and techniquesknown to the skilled person, for example by drying, passage over adessicant, flash, decantation, etc.

The effluent from water elimination step b) constitutes at least partand preferably the whole of the feed for hydrogenation step c) of theprocess of the invention.

Step c)

Step c) of the process of the invention is a step for hydrogenation ofthe unsaturated olefinic type compounds of at least a portion andpreferably the whole of the effluent derived from step b) of the processof the invention, in the presence of hydrogen and a hydrogenationcatalyst.

The effluent from step b) of the process of the invention is admitted inthe presence of hydrogen (line 8) into a hydrogenation zone 9 containinga hydrogenation catalyst which is intended to saturate the unsaturatedolefinic type compounds present in said effluent.

Preferably, the catalyst used in step c) of the invention is anon-cracking or low cracking hydrogenation catalyst comprising at leastone metal from group VIII of the periodic table of the elements andcomprising at least one support based on a refractory oxide.

Preferably, said catalyst comprises at least one metal from group VIIIselected from nickel, cobalt, ruthenium, indium, palladium and platinumand comprising at least one support based on refractory oxide selectedfrom alumina and silica-alumina.

Preferably, the metal from group VIII is selected from nickel, palladiumand platinum; highly preferably, it is selected from palladium andplatinum.

In accordance with a preferred implementation of step c) of the processof the invention, the metal from group VIII is selected from palladiumand/or platinum and the quantity of this metal is advantageously in therange 0.1% to 5% by weight, preferably in the range 0.2% to 0.6% byweight with respect to the total catalyst weight.

In accordance with a highly preferred implementation of step c) of theprocess of the invention, the metal from group VIII is palladium.

According to another preferred implementation of step c) of the processof the invention, the metal from group VIII is nickel and the quantityof this metal is advantageously in the range 5% to 25% by weight,preferably in the range 7% to 20% by weight with respect to the totalcatalyst weight.

The support for the catalyst used in step c) of the process of theinvention is a support based on a refractory oxide, preferably selectedfrom alumina and silica-alumina, more preferably alumina.

When the support is an alumina, it has a BET specific surface area whichcan limit polymerization reactions at the surface of the hydrogenationcatalyst, said surface area being in the range 5 to 140 m²/g.

When the support is a silica-alumina, the support contains a percentageof silica in the range 5% to 95% by weight, preferably in the range 10%to 80%, more preferably in the range 20% to 60% by weight and highlypreferably in the range 30% to 50%, a BET specific surface area in therange 100 to 550 m²/g, preferably in the range 150 to 500 m²/g, morepreferably less than 350 m²/g and still more preferably less than 250m²/g.

The hydrogenation step c) of the process of the invention is preferablycarried out in one or more fixed bed reactors.

In hydrogenation zone 9, the feed is brought into contact with thehydrogenation catalyst in the presence of hydrogen and at operatingtemperatures and pressures which allow hydrogenation of the unsaturatedolefinic type compounds present in the feed.

The operating conditions for hydrogenation step c) of the process of theinvention are advantageously as follows: the temperature in saidhydrogenation zone 9 is in the range 100° C. to 180° C., preferably inthe range 120° C. to 165° C., the total pressure is in the range 0.5 to6 MPa, preferably in the range 1 to 5 MPa and more preferably in therange 2 to 5 MPa. The flow rate of the feed is such that the hourlyspace velocity (ratio of the hourly volume flow rate at 15° C. for freshliquid feed to the volume of charged catalyst) is in the range 1 to 50h⁻, preferably in the range 2 to 20 h⁻¹ and more preferably in the range4 to 20 h⁻¹. The hydrogen which supplies the hydrotreatment zone isintroduced at a flow rate such that the hydrogen/hydrocarbon volumeratio is in the range 5 to 300 Nl/l/h, preferably in the range 5 to 200,more preferably in the range 10 to 150 Nl/l/h, and still more preferablyin the range 10 to 50 Nl/l/h.

Under these conditions, the unsaturated olefinic type compounds are morethan 50%, preferably more than 75% and more preferably more than 85%hydrogenated.

The effluent from step c) optionally undergoes a step for elimination ofat least a portion of the water formed during hydrogenation step c),preferably all of the water which is formed.

This water may also contain a fraction of dissolved CO and CO₂originating from the Fischer-Tropsch synthesis. This step foreliminating water takes place in the separator 11 and water iseliminated via the line 12.

It may also be advantageous to add to said step for elimination of atleast a portion of the water a nitrogen (N₂) or hydrogen (H₂) type gas(line 10) to eliminate more dissolved CO and CO₂ by stripping.

This step can also eliminate light ether type products formed during thereaction of alcohols on themselves.

The water may be eliminated using any of the methods and techniquesknown to the skilled person, for example drying, passage over adessicant, flash, decantation, etc.

At the end of step c) of the process of the invention, at least aportion and preferably all of the liquid hydrogenated effluent is sentto a hydrocracking/hydroisomerization zone 14.

Step d)

In accordance with step d) of the process of the invention, at least aportion and preferably all of the liquid hydrogenated effluent fromhydrogenation step c) of the process of the invention is sent to thehydroisomerization/hydrocracking zone 14 containing thehydroisomerization/hydrocracking catalyst, preferably at the same timeas a stream of hydrogen.

The operating conditions in which hydroisomerization/hydrocracking stepd) of the process of the invention is carried out are preferably asfollows:

The pressure is generally maintained between 0.2 and 15 MPa, preferablyin the range 0.5 to 10 MPa and advantageously in the range 1 to 9 MPa;the hourly space velocity is generally in the range 0.1 h⁻¹ to 10 h⁻¹,preferably in the range 0.2 to 7 h⁻¹ and advantageously in the range 0.5to 5.0 h⁻¹. The hydrogen ratio is generally in the range 100 to 2000normal litres of hydrogen per litre of feed per hour, preferably in therange 150 to 1500 normal litres of hydrogen per litre of feed per hour.

The temperature used in this step is generally in the range 200° C. to450° C., preferably in the range 250° C. to 450° C., advantageously inthe range 300° C. to 450° C., and more advantageously more than 320° C.or, for example, in the range 320° C. to 420° C.

Hydroisomerization and hydrocracking step d) of the process of theinvention is advantageously carried out under conditions such that theconversion per pass of products with a boiling point of 370° C. or moreinto products with boiling points of less than 370° C. is more than 80%by weight, and more preferably at least 85%, preferably more than 88%,in order to obtain middle distillates (gas oil and kerosene) withsufficiently good cold properties (pour point, freezing point) so thatthey satisfy specifications in force for this type of fuel.

The Hydroisomerization/Hydrocracking Catalysts

The majority of the catalysts in current use in hydroisomerization arebi-functional in type, associating an acid function with a hydrogenatingfunction. The acid function is supplied by supports with large surfaceareas (generally of 150 to 800 m²/g) and with a superficial acidity,such as halogenated aluminas (in particular chlorinated or fluorinated),phosphorus-containing aluminas, combinations of oxides of boron andaluminium, and silica-aluminas. The hydrogenating function is generallysupplied either by one or more metals from group VIII of the periodictable of the elements such as iron, cobalt, nickel, ruthenium, rhodium,palladium, osmium, iridium or platinum, or by a combination of at leastone metal from group VI, such as chromium, molybdenum or tungsten, andat least one group VIII metal.

In the case of bi-functional catalysts, the balance between the twofunctions, acid and hydrogenating, is the fundamental parameter whichgoverns the activity and selectivity of the catalyst. A weak acidfunction and a strong hydrogenating function produces less activecatalysts which are also less selective as regards isomerization, whilea strong acid function and a weak hydrogenating function producescatalysts which are highly active and selective as regards cracking. Athird possibility is to use a strong acid function and a stronghydrogenating function to obtain a catalyst which is highly active butalso highly selective as regards isomerization. Thus, by carefullyselecting each of the functions, it is possible to adjust theactivity/selectivity balance of the catalyst.

Advantageously, the hydroisomerization/hydrocracking catalysts arebi-functional catalysts comprising an amorphous acid support (preferablya silica-alumina) and a metallic hydro-dehydrogenating function which ispreferably provided by at least one noble metal. The support is termedamorphous, i.e. free of molecular sieve and in particular zeolite, as isthe catalyst. The amorphous acid support is advantageously asilica-alumina, but other supports may be used. When it is asilica-alumina, the catalyst preferably contains no added halogen otherthan that which may be introduced for impregnation with the noble metal,for example.

More generally and preferably, the catalyst contains no added halogen,for example fluorine. In general and preferably, the support has notundergone impregnation with a silicon compound.

In accordance with a first preferred implementation, thehydroisomerization/hydrocracking catalyst contains at least onehydrodehydrogenating element selected from noble group VIII metals,preferably platinum and/or palladium, and at least one amorphousrefractory oxide support, preferably silica-alumina.

A preferred hydroisomerization/hydrocracking catalyst used in step d) ofthe process of the invention comprises up to 3% by weight of a metal ofat least one hydro-dehydrogenating element selected from noble metalsfrom group VIII, preferably deposited on the support, and highlypreferably, the noble group VIII metal is platinum; and a supportcomprising (or preferably constituted by) at least one silica-alumina,said silica-alumina having the following characteristics:

-   -   a weight content of silica, SiO₂, in the range 5% to 95%,        preferably in the range 10% to 80%, more preferably in the range        20% to 60% and still more preferably in the range 30% to 50% by        weight;    -   a Na content of less than 300 ppm by weight, preferably less        than 200 ppm by weight;    -   a total pore volume in the range 0.45 to 1.2 ml/g, measured by        mercury porosimetry;    -   the porosity of said silica-alumina being as follows:        -   i) the volume of mesopores with a diameter in the range 40 Å            to 150 Å and with a mean diameter in the range 80 to 140 Å,            preferably in the range 80 to 120 Å, represents 20-80% of            the total pore volume measured by mercury porosimetry;        -   ii) the volume of macropores with a diameter of more than            500 Å, preferably in the range 1000 Å to 10000 Å, represents            20% to 80% of the total pore volume, measured by mercury            porosimetry;    -   a specific surface area in the range 100 to 550 m²/g, preferably        in the range 150 to 500 m²/g, more preferably less than 350 m²/g        and still more preferably less than 250 m²/g.

A second preferred hydroisomerization/hydrocracking catalyst used instep d) of the process of the invention comprises up to 3% by weight ofa metal of at least one hydro-dehydrogenating element selected fromnoble metals from group VIII of the periodic table of the elements, andpreferably, the noble group VIII metal is platinum; 0.01% to 5.5% byweight of oxide of a doping element selected from phosphorus, boron andsilicon; and a non-zeolitic support based on silica-alumina containing aquantity of more than 15% by weight and 95% by weight or less of silica(SiO₂), said silica-alumina having the following characteristics:

-   -   a mean pore diameter, measured by mercury porosimetry, in the        range 20 to 140 Å;    -   a total pore volume, measured by mercury porosimetry, in the        range 0.1 ml/g to 0.5 ml/g;    -   a total pore volume, measured by nitrogen porosimetry, in the        range 0.1 ml/g to 0.6 ml/g;    -   a BET specific surface area in the range 100 to 550 m²/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 140 Å, of less than 0.1 ml/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 160 Å, of less than 0.1 ml/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 200 Å, of less than 0.1 ml/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 500 Å, of less than 0.1 ml/g;    -   an X ray diffraction diagram which contains at least the        principal characteristic peaks of at least one of the transition        aluminas included in the group composed of alpha, rho, chi, eta,        gamma, kappa, theta and delta aluminas;    -   a settled catalyst packing density of more than 0.55 g/cm³.

Advantageously, the characteristics associated with the correspondingcatalyst are identical to those of the silica-alumina described above.

The two steps c) and d) of the process of the invention, hydrogenationand hydroisomerization-hydrocracking, may advantageously be carried outon the two types of catalysts in two or more different reactors and/orin the same reactor.

In accordance with a second preferred implementation, thehydroisomerization/hydrocracking catalyst contains at least onehydrodehydrogenating element selected from non-noble group VIII metalsand metals from group VIB and at least one amorphous refractory oxidesupport, preferably silica-alumina.

Preferably, the metal from group VIII is selected from nickel andcobalt, and the metal from group VIB is selected from molybdenum andtungsten.

Preferably, said catalyst is in the sulphide form.

A third preferred hydroisomerization/hydrocracking catalyst used in stepd) of the process of the invention comprises at least onehydro-dehydrogenating element selected from non-noble metals from groupVIII and metals from group VIB of the periodic table of the elements,preferably between 2.5% and 5% by weight of oxide of the non-nobleelement from group VIII and between 20% and 35% by weight of oxide of agroup VIB element with respect to the weight of the final catalyst;preferably, the non-noble group VIII metal is nickel and the group VIBmetal is tungsten; optionally 0.01% to 5.5% by weight of oxide of adoping element selected from phosphorus, boron and silicon; preferably,0.01% to 2.5% by weight of oxide of a doping element and a non-zeoliticsupport based on silica-alumina containing a quantity of more than 15%by weight and 95% by weight or less of silica (SiO₂), preferably aquantity of more than 15% by weight and 50% by weight or less of silica,said silica-alumina having the following characteristics:

-   -   a mean pore diameter, measured by mercury porosimetry, in the        range 20 to 140 Å;    -   a total pore volume, measured by mercury porosimetry, in the        range 0.1 ml/g to 0.5 ml/g;    -   a total pore volume, measured by nitrogen porosimetry, in the        range 0.1 ml/g to 0.6 ml/g;    -   a BET specific surface area in the range 100 to 550 m²/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 140 Å, of less than 0.1 ml/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 160 Å, of less than 0.1 ml/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 200 Å, of less than 0.1 ml/g;    -   a pore volume, measured by mercury porosimetry, included in        pores with a diameter of more than 500 Å, of less than 0.1 ml/g;    -   an X ray diffraction diagram which contains at least the        principal characteristic peaks of at least one of the transition        aluminas included in the group composed of alpha, rho, chi, eta,        gamma, kappa, theta and delta aluminas;    -   a settled catalyst packing density of more than 0.55 g/cm³.

Advantageously, the characteristics associated with the correspondingcatalyst are identical to those of the silica-alumina described above.

When the third preferred hydroisomerization/hydrocracking catalyst isused in step d) of the process of the invention, the catalyst issulphurized.

In accordance with a first preferred implementation of the process ofthe invention, in hydrogenation step c), a catalyst is used whichcontains palladium and in hydroisomerization/hydrocracking step d), acatalyst containing platinum is used.

In accordance with a second preferred implementation of the process ofthe invention, in hydrogenation step c) a catalyst containing palladiumis used and in hydroisomerization/hydrocracking step d), a sulphurizedcatalyst containing at least one hydro-dehydrogenating element selectedfrom non-noble metals from group VIII and group VIB metals is used.

In a third preferred implementation of the process of the invention, inhydrogenation step c) a catalyst containing at least one non-noblehydro-dehydrogenating element from group VIII is used and inhydroisomerization/hydrocracking step d), a sulphurized catalystcontaining at least one hydro-dehydrogenating element selected fromnon-noble group VIII metals and group VIB metals is used.

Step e)

In accordance with step e) of the process of the invention, the effluentderived from step d) undergoes the separation of unreacted hydrogen andlight gases in a gas/liquid separator 15 then recycling of the unreactedhydrogen and said light gases to the hydroisomerization/hydrocrackingstep d) (line 17).

Said light gases include light C1-C4 gases, carbon monoxide (CO), carbondioxide (CO₂) and water in the vapour form.

Said gases are separated from the liquid effluent in one or more flashdrums 15, i.e. one or more drums which carry out separation of the gasesand the liquids introduced via the line 14 b, at staged temperatures andpressures in order to increase the recovery of hydrogen. This flashstaging may advantageously be accompanied by a heat exchanger aimed atrecovering heat energy and/or cooling the effluents from the separatordrums in order to minimize losses of hydrogen.

By using an ion exchange resin upstream of the hydrotreatment andhydrocracking steps, the process of the invention can reduce the totaloxygen content of the feed and thus limit the formation of carbonmonoxide (CO) originating from the decomposition of oxygen-containingcompounds present in the feed in the hydroisomerization/hydrocrackingsection. Carbon monoxide (CO) is an inhibitor of the metallic compoundspresent on the hydroisomerization/hydrocracking catalyst, and itscontent must be minimized in order not to require an increase intemperature in order to compensate for the low activity and maintainconversion.

However, when the gaseous effluent 17 from said separation has a high COfraction, i.e. more than 10 ppm by volume, said gaseous effluent isadvantageously sent to a methanation reactor 18 in which the conversionof carbon monoxide (CO) and hydrogen into methane is advantageouslycarried out, with the aim of limiting the CO content. The principle ofmethanation, and the catalysts used are known to the skilled person andtheir use in purifying effluents containing H₂ and CO is known.

A purge is advantageously carried out (line 20) in order to eliminatethe products formed during the methanation step 18. A makeup of hydrogen(line 21) is then advantageously carried out in order to compensate forthat purge.

Step f)

In accordance with step f) of the process of the invention, the effluentfrom step e) for separating hydrogen and the light gases of the processof the invention is sent, to a distillation train 22 via a line 16,which combines atmospheric distillation with optional vacuumdistillation, which is intended to separate conversion products with aboiling point of less than 340° C. and preferably less than 370° C. andin particular including those formed during step d) in thehydroisomerization/hydrocracking reactor 14, and to separate theresidual fraction with an initial boiling point which is generally morethan at least 340° C. and preferably at least 370° C. or higher. Of theconverted and hydroisomerized products, in addition to the light C1-C4gases (line 23), at least one gasoline (or naphtha) fraction isseparated (line 24), and at least one kerosene middle distillatefraction (line 25) and a gas oil fraction (line 27) are separated.Preferably, the residual fraction, with an initial boiling point whichis generally over at least 340° C. and preferably at least 370° C. isrecycled (line 28) to step d) of the process of the invention to thehead of the hydroisomerization and hydrocracking zone 14. In accordancewith another implementation of step f) of the process of the invention,said residual fraction may supply excellent oil bases.

It may also be advantageous to recycle (line 26) at least part andpreferably all of at least one of the kerosene and gas oil cuts obtainedto step d) (line 14). The gas oil and kerosene cuts are preferablyrecovered separately or mixed, but the cut points are adjusted by theoperator as a function of requirements. It has been shown that it isadvantageous to recycle part of the kerosene to improve its coldproperties.

Products Obtained

The gas oil(s) obtained have a pour point of at most 0° C., generallyless than −10° C. and usually less than −15° C. The cetane index is morethan 60, generally more than 65, and usually more than 70.

The kerosene(s) obtained have a freezing point of at most −35° C.,generally less than −40° C. The smoke point is more than 25 mm,generally more than 30 mm In this process, gasoline (unwanted)production is as low as possible. The gasoline yield is always less than50% by weight, preferably less than 40% by weight, advantageously lessthan 30% by weight or 20% by weight or even 15% by weight.

EXAMPLE 1 Implementation of the Process of the Invention

The C5+ paraffinic effluent derived from a Fischer-Tropsch synthesisunit is described in Table 1.

TABLE 1 Composition of C5+ fraction of FT effluent Units Effluent fromFT unit Density @ 15° C. — 0.782 Paraffins content wt % 80 Olefinscontent wt % 15 Alcohols content wt % 3 Acid content wt % 2 Estercontent wt % 2 CO content ppm by weight 30 CO₂ content ppm by weight 390Organometallics ppm 5 Simulated distillation Initial boiling point ° C.25  5% by weight ° C. 50 10% by weight ° C. 77 30% by weight ° C. 20050% by weight ° C. 300 70% by weight ° C. 400 90% by weight ° C. 530 95%by weight ° C. 575 End point ° C. 650 370° C. + fraction wt % 35 Waterppm by weight 276

Step a)

The C5+ fraction 1 passed over an ion exchange resin with trade nameAmberlyst 35 sold by Röhm & Haas, said resin allowing simultaneouscapture of metals dissolved in the feed and esterification of alcoholsand carboxylic acids to esters. Said resin was constituted by divinylbenzene—polystyrene copolymers with a degree of cross-linking of 20 andan acid strength, assayed by potentiometry during neutralization with aKOH solution, of 4.15 mmol H+ equivalent/g.

Step a) was carried out at a temperature of 110° C., a pressure of 1MPa, and at an hourly space velocity of 1 h⁻¹. Under these conditions,95% of the acids were esterified, the analysis of the conversion of theacids being given by the difference in potassium hydroxide titrationbetween the feed and the effluent using the ASTM D664 method. Thecomposition of the outlet effluent is given in Table 2.

TABLE 2 Composition of effluent derived from step a) afteresterification Units Effluent from FT unit Density @ 15° C. — 0.782Paraffins content wt % 80 Olefins content wt % 15 Alcohols content wt %1 Acid content wt % <0.1 Ester content wt % 4 CO content ppm by weight30 CO₂ content ppm by weight 390 Organometallics Ppm <1 Simulateddistillation Initial boiling point ° C. 25  5% by weight ° C. 50 10% byweight ° C. 77 30% by weight ° C. 200 50% by weight ° C. 300 70% byweight ° C. 400 90% by weight ° C. 530 95% by weight ° C. 575 End point° C. 650 370° C. + fraction wt % 35 Water ppm by weight 2500

Step b)

The effluent derived from step a) then underwent the elimination of thewater formed during said step a), by decanting/coalescence in suitableequipment known to the skilled person.

Step c)

All of the effluent derived from water elimination step b) thenunderwent a step for hydrogenation in the presence of hydrogen and ahydrogenation catalyst with trade name LD265 sold by Axens, saidcatalyst comprising 0.3% by weight of palladium deposited on an aluminawith a specific surface area of 69 m²/g.

Hydrogenation step c) was carried out at a reaction temperature of 130°C., at a pressure of 3.5 MPa, the hydrogen was introduced at a flow ratesuch that the hydrogen/hydrocarbon volume ratio was 32 Nl/l/h, and withan hourly space velocity of 10 h⁻¹. Under these conditions, theconversion of olefins was 85% by weight.

The liquid effluent derived from hydrogenation step c) had thecomposition described in Table 3:

TABLE 3 Composition of effluent derived from step c) Units Effluent fromFT unit Density @ 15° C. — 0.782 Paraffins content wt % 93 Olefinscontent wt % 2 Alcohols content wt % 1 Acid content wt % <0.1 Estercontent wt % 4 CO content ppm by weight 0 CO₂ content ppm by weight 390Organometallics Ppm <1 Simulated distillation Initial boiling point ° C.25  5% by weight ° C. 50 10% by weight ° C. 77 30% by weight ° C. 20050% by weight ° C. 300 70% by weight ° C. 400 90% by weight ° C. 530 95%by weight ° C. 575 End point ° C. 650 370° C. + fraction wt % 35 Waterppm by weight 300

Step d)

All of the effluent from hydrogenation step c) underwent ahydroisomerization/hydrocracking step in the presence of fresh hydrogenand a hydroisomerization/hydrocracking catalyst, in which the residualfraction with an initial boiling point of more than 370° C., unreactedhydrogen and light gases were recycled.

The hydroisomerization/hydrocracking catalyst comprised 0.6% by weightof platinum and a support comprising 29.3% by weight of silica, SiO₂,and 70.7% by weight of alumina, Al₂O₃, a Na content of 100 ppm byweight, a total pore volume comprising 0.69 ml/g measured by mercuryporosimetry, a volume of mesopores with a mean diameter of 80 Årepresenting 78% of the total pore volume measured by mercuryporosimetry, a volume of macropores with a diameter of more than 500 Årepresenting 22% of the total pore volume measured by mercuryporosimetry and a specific surface area of 300 m²/g.

The hydroisomerization/hydrocracking step was carried out under theconditions described in Table 4.

The conversion per pass for products with a boiling point of 370° C. ormore into products with a boiling point of less than 370° C. was 85%.

TABLE 4 Operating conditions for hydroisomerization/hydrocracking stepUnit H₂ partial pressure MPa 5 Space velocity, HSV h⁻¹ 1.0 Reactiontemperature ° C. 345 Hydrogen ratio N1/1 600

Step e)

The effluent from the hydroisomerization/hydrocracking step underwentseparation, in a gas/liquid separator, of the unreacted hydrogen andlight gases which were recycled to the hydroisomerization/hydrocrackingstep in order to recover a liquid effluent. The carbon monoxide (CO)content generated per pass in the gaseous effluent was limited to 1.1%by weight.

Step f)

The liquid effluent from the separation step e) was then sent to adistillation train to separate the light products formed during thesesteps: the gases (C1-C4), a gasoline cut, a gas oil cut and a kerosenecut, and also a fraction, termed the residual fraction, which had aninitial boiling point equal to 370° C. which was recycled in itsentirety to the inlet to the hydroisomerization/hydrocracking reactor inorder to maximize the production of gas oil and kerosene.

The yields are given in Table 5.

TABLE 5 Yield of various cuts after separation Wt % Boiling point C1-C41.9 −161° C. to 35° C.  Naphtha 12.1  35° C. to 150° C. Kerosene 34.5150° C. to 250° C. Gas oil 51.6 250° C. to 370° C.

Example 2 Comparative

A process for producing middle distillates from a C5+ liquid paraffinicfraction, termed a heavy fraction, with an initial boiling point in therange 15° C. to 40° C. produced by Fischer-Tropsch synthesis which wasthe same as that used in Example 1 was carried out, comprising a stepfor hydrogenation followed by a hydroisomerization/hydrocracking step,with no prior step for passage over at least one ion exchange resin;this was carried out for comparison purposes.

The C5+ paraffinic fraction derived from the Fischer-Tropsch synthesisunit was described in Table 1 of Example 1.

Hydrogenation Step

The C5+ liquid paraffinic fraction underwent a step for hydrogenation inthe presence of hydrogen and a hydrogenation catalyst with trade nameLD265 sold by Axens, said catalyst comprising 0.3% by weight ofpalladium deposited on an alumina with a specific surface area of 69m²/g.

In order to maintain a conversion into olefins of 85% by weight, as forExample 1, the hydrogenation step was carried out at a reactiontemperature of 150° C., at a pressure of 3.5 MPa, the hydrogen wasintroduced at a flow rate such that the hydrogen/hydrocarbon volumeratio was 32 Nl/l/h and at an hourly space velocity of 8 h⁻¹.

Under these conditions, the conversion into olefins was maintained at85% by weight.

The liquid effluent derived from hydrogenation step c) had thecomposition described in Table 9:

TABLE 9 Composition of effluent derived from the hydrogenation stepUnits Effluent from FT unit Density @ 15° C. — 0.782 Paraffins contentwt % 91 Olefins content wt % 2 Alcohols content wt % 3 Acid content wt %2 Ester content wt % 2 CO content ppm by weight 30 CO₂ content ppm byweight 390 Organometallics ppm <1 Simulated distillation Initial boilingpoint ° C. 25  5% by weight ° C. 50 10% by weight ° C. 77 30% by weight° C. 200 50% by weight ° C. 300 70% by weight ° C. 400 90% by weight °C. 530 95% by weight ° C. 575 End point ° C. 650 370° C. + fraction wt %35 Water ppm by weight 276

Hydroisomerization/Hydrocracking Step

All of the effluent from the hydrogenation step underwent ahydroisomerization/hydrocracking step in the presence of fresh hydrogenand a hydroisomerization/hydrocracking catalyst, in which the residualfraction with an initial boiling point of more than 370° C., unreactedhydrogen and light gases were recycled.

The hydroisomerization/hydrocracking catalyst comprised 0.6% by weightof platinum and a support comprising 29.3% by weight of silica, SiO₂,and 70.7% by weight of alumina, Al₂O₃, a Na content of 100 ppm byweight, a total pore volume comprising 0.69 ml/g measured by mercuryporosimetry, a volume of mesopores with a mean diameter of 80 Årepresenting 78% of the total pore volume measured by mercuryporosimetry, a volume of macropores with a diameter of more than 500 Årepresenting 22% of the total pore volume measured by mercuryporosimetry and a specific surface area of 300 m²/g.

The hydroisomerization/hydrocracking step was carried out under theconditions described in Table 10.

To maintain the conversion per pass for products with a boiling point of370° C. or more into products with a boiling point of less than 370° C.at 85%, the temperature was increased and adjusted to 370° C.

TABLE 10 Operating conditions for hydroisomerization/hydrocracking stepUnit H₂ partial pressure MPa 5 Space velocity, HSV h⁻¹ 1.0 Reactiontemperature ° C. 370 Hydrogen ratio N1/1 600

Separation Step

The effluent from the hydroisomerization/hydrocracking step underwentseparation, in a gas/liquid separator, of the unreacted hydrogen andlight gases which were recycled to the hydroisomerization/hydrocrackingstep in order to recover a liquid effluent. The carbon monoxide (CO)content generated per pass in the gaseous effluent was limited to 2% byweight.

The presence of carbon monoxide (CO) derived from the decomposition ofoxygen-containing compounds in the hydrocracking section not eliminatedby passage over at least one ion exchange resin prior to hydrogenationand to hydroisomerization/hydrocracking, and being an inhibitor of theactivity of the hydroisomerization/hydrocracking catalyst activity,necessitated an increase in the reaction temperature in order tomaintain the conversion.

Final Distillation Step

The liquid effluent from the separation step was then sent to adistillation train to separate the light products formed during thesesteps: the (C1-C4) gases, a gasoline cut, a gas oil cut and a kerosenecut, and also a fraction, termed the residual fraction, which had aninitial boiling point equal to 370° C. which was recycled in itsentirety to the inlet to the hydroisomerization/hydrocracking reactor inorder to maximize the production of gas oil and kerosene.

The yields are given in Table 11.

TABLE 11 Yield of various cuts after separation Wt % Boiling point C₁-C₄2.6 −161° C. to 35° C.  Naphtha 15  35° C. to 150° C. Kerosene 35 150°C. to 250° C. Gas oil 47.4 250° C. to 370° C.

The presence of carbon monoxide (CO) in the gaseous effluent which wasan inhibitor of the hydrogenating function of the hydrocracking catalystmodified not only the activity but also the selectivity of said catalystfor middle distillates.

It can be seen that the absence of a prior step for passage over atleast one ion exchange resin, necessitating an increase in temperaturein order to maintain the conversion, resulted in an increase in theproduction of an unwanted light gas and naphtha fraction byover-cracking.

Thus, by using an ion exchange resin upstream of the hydrotreatment andhydroisomerization/hydrocracking steps, the process of the inventionexemplified in Example 1 can reduce the total oxygen content of the feedand thereby limit the formation of carbon monoxide (CO) originating fromthe decomposition of oxygen-containing compounds present in the feed inthe hydroisomerization/hydrocracking section. In fact, carbon monoxide(CO) is an inhibitor of the metallic compounds present on thehydroisomerization/hydrocracking catalyst and its content must beminimized in order not to require an increase in temperature in order tocompensate for the drop in activity and to maintain conversion.

Thus, it can be seen that the process of the invention can reduce theproduction of carbon monoxide (CO) (1.1% by weight) by carrying out stepa) compared with a process which is not in accordance with the inventionwhich does not carry out said step a) for passage over at least one ionexchange resin, and can reduce the temperatures employed in thehydrogenation and hydroisomerization/hydrocracking steps in order toobtain the same conversion of 85% of products with a boiling point of370° C. or more into products with a boiling point of less than 370° C.

1. A process for producing middle distillates from a C5+ liquidparaffinic fraction, termed a heavy fraction, with an initial boilingpoint in the range 15° C. to 40° C., produced by Fischer-Tropschsynthesis, comprising the following steps in succession: a) passing saidC5+ liquid paraffinic fraction, termed a heavy fraction, over at leastone ion exchange resin to allow esterification of alcohols andcarboxylic acids into esters and/or to retain metals dissolved in thefeed, at a temperature in the range 80° C. to 150° C., at a totalpressure in the range 0.7 to 2.5 MPa, at an hourly space velocity in therange 0.2 to 2.5 h⁻¹; b) eliminating at least a portion of the waterformed in step a); c) hydrogenating the unsaturated olefinic typecompounds of at least a portion of the effluent derived from step b) inthe presence of hydrogen and a hydrogenation catalyst; d)hydroisomerization/hydrocracking of at least a portion of thehydrotreated effluent derived from step c) in the presence of hydrogenand a hydroisomerization/hydrocracking catalyst; e) separating andrecycling unreacted hydrogen and light gases to thehydroisomerization/hydrocracking step d); f) distilling the effluentderived from step e).
 2. A process according to claim 1, in which saidC5+ liquid fraction undergoes a step for decontamination by passage overa guard bed containing at least one guard bed catalyst, before passageover an ion exchange resin in accordance with step a).
 3. A processaccording to claim 2, in which said guard bed catalyst comprises amacroporous mercury volume for a mean diameter at 50 nm of more than 0.1cm³/g, and a total volume of more than 0.60 cm³/g.
 4. A processaccording to claim 1, in which said C5+ liquid paraffinic fractionpasses over a single ion exchange resin in order to carry out thesimultaneous esterification of alcohols and carboxylic acids into estersand capture of metals dissolved in the feed.
 5. A process according toclaim 4, in which said resin is used at a temperature in the range 100°C. to 150° C., at a pressure in the range 1 to 2 MPa and at an hourlyspace velocity in the range 0.5 to 1.5 h⁻¹.
 6. A process according toclaim 4, in which said resin is constituted by copolymers of divinylbenzene and polystyrene with a degree of cross-linking in the range 20%to 35%, and an acid strength, assayed by potentiometry duringneutralization with a KOH solution, in the range 0.2 to 6 mmol H+equivalent/g.
 7. A process according to claim 4, in which said resin isa polysiloxane grafted with alkylsulphonic type acid groups (of the—CH₂—CH₂—CH₂—SO₃H type), with a size in the range 0.5 to 1 2 mm and withan acid strength, assayed by potentiometry during neutralization with aKOH solution, of 0.4 to 1.5 mmol H+ equivalent/g.
 8. A process accordingto claim 1, in which said C5+ liquid paraffinic fraction passes over twodistinct ion exchange resins with different natures, in two differentreactors.
 9. A process according to claim 8, in which the reactorcontaining the ion exchange resin allowing the capture of metals is usedupstream of the reactor containing the ion exchange resin allowing theesterification of alcohols and carboxylic acids.
 10. A process accordingto claim 8, in which said first resin is a resin constituted bycopolymers of divinyl benzene and polystyrene with a degree ofcross-linking in the range 1% to 20% and an acid strength, assayed bypotentiometry during neutralization with a KOH solution, in the range 1to 15 mmol H+ equivalent/g.
 11. A process according to claim 8, in whichsaid first resin is used at a temperature in the range 80° C. to 110°C., at a pressure in the range 1 to 2 MPa and at an hourly spacevelocity in the range 0.2 to 1.5 ⁻¹.
 12. A process according to claim 8,in which the hydroisomerization/hydrocracking catalyst contains at leastone hydrodehydrogenating element selected from noble metals from groupVIII, preferably platinum and/or palladium, and at least one amorphousrefractory support, preferably silica-alumina.
 13. A process accordingto claim 1, in which the paraffinic feed produced by Fischer-Tropschsynthesis is produced from a synthesis gas produced from a natural gasusing the gas-to-liquid, GTL, route.
 14. A process according to claim 1,in which the paraffinic feed produced by Fischer-Tropsch synthesis isproduced from a synthesis gas produced from coal using thecoal-to-liquid, CTL, route.
 15. A process according to claim 1, in whichthe paraffinic feed produced by Fischer-Tropsch synthesis is producedfrom a synthesis gas produced from biomass using the biomass-to-liquidroute.